JP2024500349A - Hydroformylation reaction method - Google Patents

Abstract

The present invention relates to a hydroformylation reaction method. In one aspect, a hydroformylation reaction method comprises: (a) contacting an olefin with hydrogen and carbon monoxide in the presence of a homogeneous catalyst in a reactor to provide a reaction fluid, the reactor comprising one or more reactors. (b) removing a portion of the reaction fluid from the first reaction zone; and (c) passing at least a portion of the removed reaction fluid through a shear mixing device to remove the removed reaction fluid. (d) creating bubbles in a portion of the reaction fluid, wherein at least a portion of the hydrogen and carbon monoxide provided to the reactor is introduced through a shear mixing device; and (d) the removed reaction fluid. through one or more nozzles to the first reaction zone, the returned reaction fluid exiting each nozzle being a jet, the mixing energy density provided to the reactor by the jets; is 500 watts/m3 or more.

Description

TECHNICAL FIELD This invention relates generally to hydroformylation reaction methods.

Introduction Hydroformylation is the reaction of olefins with _H2 and CO in the presence of a homogeneous rhodium catalyst modified with organophosphorus ligands to produce aldehydes according to the formula below.

Typically, hydroformylation reactions are carried out in a liquid phase in which synthesis gas (a gas mixture of _H2 and CO) is sparged into a reaction fluid containing liquid olefins, product aldehydes, heavy materials, and a homogeneous rhodium/ligand catalyst. It will be held in

Effective gas-liquid mixing is important in both initiation and maintenance of the hydroformylation reaction, as H ₂ and CO must be dissolved in the reaction fluid for the reaction to occur.

In addition, it is necessary to remove the heat generated by the exothermic hydroformylation reaction and to control the reactor temperature at the desired reaction conditions. This is typically accomplished by internal cooling coils or by recycling the reaction fluid through an external heat exchanger and returning the cooled reaction fluid to the reactor, or both. Ru.

Furthermore, under the same conditions as the hydroformylation reaction described above, the aldehyde obtained may be further reacted and hydrogenated in situ to yield the corresponding alcohol. Hydroformylation under amination conditions can then be considered a variant of the hydroformylation reaction.

Another secondary catalytic activity of some hydroformylation catalysts is the hydrogenation or isomerization of double bonds (e.g., olefins with internal double bonds) to saturated hydrocarbons or α-olefins, and vice versa. . In order to establish and maintain specific hydroformylation reaction conditions within the reactor, it is important to avoid secondary reactions of these hydroformylation catalysts. Even small deviations from the process parameters can lead to the formation of significant amounts of undesired secondary products, thus ensuring that virtually identical Maintaining process parameters can be very important. Furthermore, the volume within the reactor in which the syngas is not sufficiently dispersed or dissolved does not contribute to either the reaction or the productivity of the reactor. In addition, many hydrolyzable catalysts exhibit catalyst degradation in the absence of syngas at reaction temperatures, and regions of low dispersion or dissolved syngas contribute to reduced catalyst performance. Alternatively, areas of excessive dissolved syngas concentration also need to be avoided, as many rhodium phosphine catalysts exhibit decomposition in high CO environments. Therefore, highly dispersed and uniform gas mixing (as determined by high gas stagnation or gas fraction) is the most desirable outcome. Generally speaking, in the hydroformylation of olefins over homogeneous rhodium catalysts modified with organophosphorous ligands, it is advantageous to establish the optimal concentrations of hydrogen and carbon monoxide dissolved in the liquid reaction medium.

The concentration of dissolved carbon monoxide (CO) in the reaction solution is of particular importance and is an important control variable in hydroformylation reactors. Dissolved CO concentration in the reaction liquid cannot be measured directly, but is typically monitored and approximated using on-line analytical equipment, and is estimated to be in equilibrium with the reaction liquid phase in the reactor vapor space. Measure the CO partial pressure. This approximation is improved if the reaction fluids in the reactor are more uniformly mixed to better approximate a fully backed-mixed reaction mixture, such as the classical CSTR model.

To achieve high conversion rates, reactors with multiple zones, such as those described in US Pat. No. 5,728,893, are preferred. However, in reactors with two or more reaction zones, measuring the CO partial pressure in the headspace only indicates the CO concentration in the upper zone and not necessarily the CO concentration in the lower reaction zone. Not exclusively. This becomes more important if the upper reaction zone is not a back-mixed reactor. In the latter case, it is even more important that the feed to the non-backmixing reaction zone is as homogeneous as possible in order to achieve as uniform and/or predictable a CO distribution as possible.

The hydrocarbon (paraffin) forming reactions, the formation of high boiling condensates of aldehydes (ie, high boilers or "heavy products"), and the rate of decomposition of organophosphorus-rhodium based catalysts are also affected by the reaction temperature. In the case of backmix reactors, it is important to avoid the formation of gradients with respect to the reaction temperature and the concentration of dissolved CO in the volume of reaction liquid present in the reactor. In other words, it is important that substantially identical operating conditions be established and maintained throughout the entire liquid volume. Therefore, it is preferable to avoid non-uniform distribution of reactants and temperature within the reaction zone. However, other non-backmixing reactors such as plug flow reactors and bubble reactors are known to have gradients and therefore it is more difficult to control these reactors in this way.

Means of supplying synthesis gas and ensuring good distribution have long been recognized. Academic papers have focused on stirring speed, for example, and the technology disclosed in WO 2018/236823 for back-mixing reactors without mechanical stirrers has shown that a good distribution of synthesis gas is possible. The study teaches that the process is important for high reactivity and reactor performance.

Therefore, a hydroformylation reactor design, preferably multi-zone, that provides a highly dispersed and uniform syngas and temperature distribution within the reactor without the use of mechanical agitators and establishes good initial syngas distribution. It is desirable to have a hydroformylation reactor design.

The present invention generally relates to a hydroformylation reaction process for preparing aldehydes by reacting olefins with carbon monoxide and hydrogen gas in the liquid phase. In the presence of a catalyst, in various embodiments at elevated temperatures from 50° C. to 145° C. and pressures from 1 to 100 bar, a portion of these gases is dispersed in the reaction liquid in the form of bubbles, and another portion is dispersed in the reaction liquid. dissolved in Embodiments of the present invention can advantageously provide thorough gas-liquid mixing of reaction fluids within the reactor without the use of mechanical agitators.

Using high-velocity fluid flow, the reactor contents can be mixed simply by (1) introducing the synthesis gas as a well-distributed stream of microbubbles and (2) imparting momentum and shear to the reaction liquid. It has been found that by dispersing the syngas bubbles without any oxidation, the bubbles can be evenly distributed and mixed throughout the reaction zone. Although not located at the top of the reactor as in traditional Venturi gas-liquid mixing reactor designs, in some embodiments of the invention the entire reactor fluid is located higher within the reactor. Significantly uniform temperatures and gas-liquid mixing can be achieved, as evidenced by more uniform gas fractions or gas loads and constant and uniform temperatures. In addition, uniformly mixed microbubbles facilitate the introduction of process fluids into non-backmixing reaction zones such as bubble columns or plug flow reactors, which is difficult with venturi-type reactor designs.

In one aspect, a hydroformylation reaction method comprises: (a) contacting an olefin with gaseous hydrogen and carbon monoxide in the presence of a homogeneous catalyst in a reactor to provide a reaction fluid, the reactor comprising: (b) removing a portion of the reaction fluid from the first reaction zone; and (c) passing at least a portion of the removed reaction fluid through a shear mixing device to remove the reaction fluid. (d) generating gas bubbles in a portion of the reactant fluid that has been removed, wherein at least a portion of the hydrogen and carbon monoxide provided to the reactor is introduced through a shear mixing device; returning the reaction fluid to the first reaction zone through one or more nozzles, the returned reaction fluid exiting each nozzle being a jet, the mixing provided by the jet to the reactor; The energy density is given by the following formula:

(where V is the volume of reaction fluid in the first reaction zone (m ³ ) and N is the volume of reaction fluid in the first reaction zone (m 3 ), and N is the is the total number of jets returned to the first reaction zone, as identified in , and ρ _i is the average density (kg /m ³ ), Q _i is the volumetric flow rate (m 3 /s) of the reaction fluid returned to the first reaction zone through the i-th jet, and A _i is the volumetric flow rate (m ³ /s) of the reaction fluid flowing through the i-th of the nozzle at the location where the reaction fluid exits the nozzle and enters the first reaction zone (m ² ).

These and other embodiments are described in more detail in the Detailed Description.

1 is a schematic diagram showing an example of a hydroformylation reactor and related equipment that can be used in a hydroformylation reaction method according to an embodiment of the present invention. FIG. 2 is a schematic diagram illustrating the angles and other parameters at which a nozzle may be oriented within a reactor, according to some embodiments of the invention. Illustrated are two embodiments of shear mixing devices that may be used in some embodiments of the present invention, where “G” represents gas entering the shear mixing device and “L” represents liquid entering the shear mixing device. represents. 2 is a series of views showing different positions of a nozzle within a reactor, different positions of one or more donut baffles relative to the jet, and the angle of the jet exiting the nozzle, according to some embodiments of the invention. FIG. 1 shows a gas volume fraction color-coded diagram of Comparative Example A and Examples 1 and 2 of the present invention. A color-coded diagram of average values of mass transfer coefficients (KLa) of Comparative Example A and Examples 1 and 2 of the present invention is shown.

Hydroformylation processes generally involve CO, H ₂ , and at least one olefin. Optional process components include amines and/or water.

All references to the Periodic Table of Elements and the various groups therein can be found in the CRC Handbook of Chemistry and Physics, 72nd Ed. (1991-1992) to the version published in CRC Press, page I-10.

Unless stated to the contrary or implied by context, all parts and percentages are by weight and all test methods are as of the filing date of this application. As used herein, the term "ppmw" means parts per million by weight. For purposes of U.S. patent practice, the content of any referenced patent, patent application, or publication shall include, among other things, the disclosure of definitions (to the extent not inconsistent with any definitions specifically provided in this disclosure) and the art in the art. (or equivalent U.S. versions thereof) are incorporated by reference in their entirety.

As used herein, "a," "an," "the," "at least one," and "one or more" are used interchangeably. The words "comprise", "include", and variations thereof do not have a limiting meaning when these terms appear in the specification and claims. Thus, for example, an aqueous composition comprising "a" particles of a hydrophobic polymer can be interpreted to mean that the composition comprises "one or more" particles of a hydrophobic polymer.

Also, herein, the recitation of numerical ranges by endpoints includes all numbers subsumed within that range (eg, 1 to 5 is 1, 1.5, 2, 2.75, 3, 3, etc.). 80, 4, 5, etc.). For purposes of this invention, consistent with the understanding of those skilled in the art, numerical ranges are understood to be intended to include and support all subranges that may be subsumed within that range. sea bream. For example, a range of 1 to 100 is intended to convey 1.01 to 100, 1 to 99.99, 1.01 to 99.99, 40 to 60, 1 to 55, and so on.

As used herein, the term "hydroformylation" refers to, but is not limited to, one or more substituted or unsubstituted olefinic compounds or a reaction involving one or more substituted or unsubstituted olefinic compounds. Intended to include all permissible asymmetric and non-asymmetric hydroformylation processes that involve converting a mixture into one or more substituted or unsubstituted aldehydes or a reaction mixture containing one or more substituted or unsubstituted aldehydes. be done.

As used herein, the terms "reaction fluid," "reaction medium," and "catalyst solution" are used interchangeably and refer to (a) a metal-organophosphorus ligand complex catalyst, (b) a free organophosphorus ligand. , (c) aldehyde products formed in the reaction, (d) unreacted reactants (e.g., hydrogen, carbon monoxide, olefins), (e) the aforementioned metal-organophosphorus ligand complex catalyst and the aforementioned a solvent for the free organophosphorus ligand and, optionally, (f) one or more ligand decomposition products (homogeneous or heterogeneous) such as oxides and phosphoric compounds formed in the reaction. These compounds may include, but are not limited to, those deposited on the surfaces of process equipment. It should be understood that the reaction fluid can be a mixture of gas and liquid. For example, the reaction fluid may contain gas bubbles (e.g., hydrogen and/or CO and/or inert substances) entrained within the liquid or gases (e.g., hydrogen and/or CO and/or inert substances) dissolved in the liquid. may be included. The reaction fluid includes, but is not limited to, (a) fluid within the reaction zone, (b) fluid flow en route to the separation zone, (c) fluid within the separation zone, (d) recirculation flow, (e) reaction zone. or fluid withdrawn from the separation zone, (f) withdrawn fluid treated with an aqueous buffer solution, (g) treated fluid returned to the reaction zone or separation zone, (h) on its way into an external cooler. (i) fluid in the external cooler, (j) fluid returned from the external cooler to the reaction zone, and (k) ligand decomposition products and their salts.

As used herein, the term "first reaction zone" in a multiple reaction zone reactor or reaction train refers to the reaction zone where the majority of the catalyst is introduced (e.g., an upstream reaction zone that is not part of this invention). Recycled catalyst or catalyst-containing reaction fluid from the reactor). A "second reaction zone" follows the first reaction zone, and so on, in that the majority of the catalyst flows from the first reaction zone to the second reaction zone. The advantages of this type of reaction scheme are described in US Pat. No. 5,728,893. For the purposes of the present invention, the term "first reaction zone" refers to the reaction zone where the majority of the olefins, synthesis gas, and catalyst are introduced into the reactor. The majority of the reaction fluid leaving this first reaction zone is then transported through the perforated plate to the "second reaction zone" without intermediate piping. In this context, "first" and "second" relate to the path followed by the majority of the catalyst in this reactor, and that there may be a reaction zone before this reactor body which is not included in the present invention. is recognized.

The present invention generally relates to a hydroformylation reaction process for preparing aldehydes by reacting olefins with carbon monoxide and hydrogen gas in the liquid phase. Embodiments of the invention advantageously disperse at least a portion of the carbon monoxide and/or hydrogen gas in the form of small bubbles within the reaction fluid. In some embodiments, the methods of the present invention can advantageously provide thorough gas-liquid mixing of reaction fluids without the use of mechanical agitators.

In one embodiment, the hydroformylation method of the present invention comprises: (a) contacting an olefin, hydrogen, and carbon monoxide in the presence of a homogeneous catalyst in a reactor to provide a reaction fluid, the method comprising: (b) removing a portion of the reaction fluid from the first reaction zone; and (c) passing at least a portion of the removed reaction fluid through a shear mixing device. (d) generating bubbles in a portion of the removed reaction fluid, wherein at least a portion of the hydrogen and carbon monoxide provided to the reactor is introduced through a shear mixing device; returning the returned reaction fluid to the first reaction zone through one or more nozzles, the returned reaction fluid exiting each nozzle being a jet, the returned reaction fluid being provided to the reactor by the jet. The mixing energy density is given by the following formula:

(where V is the volume of reaction fluid in the first reaction zone (m ³ ) and N is the volume of reaction fluid in the first reaction zone (m 3 ), and N is the is the total number of jets returned to the first reaction zone, as identified in , and ρ _i is the average density (kg /m ³ ), Q _i is the volumetric flow rate (m 3 /s) of the reaction fluid returned to the first reaction zone through the i-th jet, and A _i is the volumetric flow rate (m ³ /s) of the reaction fluid flowing through the i-th of the nozzle at the location where the reaction fluid exits the nozzle and enters the first reaction zone (m ² ). In some embodiments, in addition to the hydrogen and carbon monoxide provided to the reactor through the shear mixer, an inert gas (e.g., methane, _CO2 , argon, nitrogen, etc.) is also provided to the reactor through the shear mixer. may be present in the synthesis gas provided to the vessel. In some embodiments, the average bubble size of the bubbles produced by the shear mixing device is between 10 nanometers and 3,000 microns. In some embodiments, the average bubble size of the bubbles produced by the shear mixing device is between 100 microns and 800 microns.

The flow rate of the reaction fluid through the shear mixing device can be important to promote proper mixing of the reaction fluid. In one embodiment, the flow rate of the reactant fluid through the shear mixing device is:
q _SM >525 (μ _o /ρ _o ) P _SM
(where q _SM is the flow rate of the reaction fluid entering the shear mixing device (m ³ /s), ρ _o is the density of the reaction fluid before entering the shear mixing device (kg/m ³ ), μ _o is the viscosity of the reaction fluid before entering the shear mixing device (Pa·s) and P _SM is the minimum wetted perimeter of the cross section of the liquid flow inside the shear mixing device.

In some embodiments, the removed reaction fluid is returned to the first reaction fluid through at least two nozzles, each nozzle having an angle (alpha) of the nozzle with respect to the horizontal plane between +75° and −75°. oriented such that alpha, the angle of the nozzle with respect to the vertical plane passing through the center of the reactor (beta), and the distance from the vertical plane passing through the center of the reactor (phi) if beta is zero are all non-zero. It is.

In some embodiments, hydrogen and carbon monoxide are provided as syngas, and at least 20% of the syngas provided to the first reaction zone passes through a shear mixing device before entering the first reaction zone. do.

In some embodiments, at least a portion of the synthesis gas is introduced into the cylindrical reactor through the sparger at a height that is less than 50% of the reaction fluid fill height of the first reaction zone.

In some embodiments, the reactor includes a horizontally oriented ring baffle attached to the interior wall of the reactor, the ring baffle being less than 90% of the height of the liquid reaction fluid in the first reaction zone. The solid portion of the ring baffle spans 5-30% of the reactor diameter.

In some embodiments, a stirrer is placed within the reactor. In some embodiments, the agitator is not operational. In some embodiments, the agitator and returned reaction fluid provide a mixing energy density within the cylindrical reactor.

In some embodiments, the reactor is vertically oriented.

The reactor, in some embodiments, further includes a second reaction zone, and the reaction fluid flows from the first reaction zone to the second reaction zone without piping. In some further embodiments, the first reaction zone and the second reaction zone are separated by a perforated plate. The reactor, in some embodiments, further includes a third reaction zone, and the reaction fluid flows from the second reaction zone to the third reaction zone without piping. In some further embodiments, the second reaction zone and the third reaction zone are separated by a perforated plate.

In some embodiments, the reactor includes a product outlet nozzle located in the lower portion of the reactor and a means for preventing gas entrainment located in the bottom volume of the reactor.

The hydroformylation process of the present invention involves contacting an olefin, hydrogen, and carbon monoxide in the presence of a homogeneous catalyst in a reactor to provide a reaction fluid, the reactor having one or more reaction zones. include, contain.

Hydrogen and carbon monoxide may be obtained from any suitable source, including petroleum cracking and refining operations. Syngas mixture is the preferred source of hydrogen and CO. Syngas (synthesis gas) is the name given to gas mixtures containing varying amounts of CO and _H2 . The generation method is well known. Hydrogen and CO are typically the main components of syngas, but syngas may contain _CO2 and inert gases such as _N2 and Ar. The molar ratio of H ₂ to CO varies widely, but generally ranges from 1:100 to 100:1, usually from 1:10 to 10:1. Synthesis gas is commercially available and is often used as a fuel source or as an intermediate for the production of other chemicals. The H ₂ :CO molar ratio for chemical production is often between 3:1 and 1:3, with about 1:2 to 2:1 typically targeted for most hydroformylation applications. Ru.

In typical embodiments of the hydroformylation process, a solvent is advantageously used. Any suitable solvent that does not unduly interfere with the hydroformylation process may be used. By way of illustration, suitable solvents for rhodium-catalyzed hydroformylation processes include, for example, U.S. Pat. , 929,289. Non-limiting examples of suitable solvents include saturated hydrocarbons (alkanes), aromatic hydrocarbons, water, ethers, aldehydes, ketones, nitriles, alcohols, esters, and aldehyde condensation products. Specific examples of solvents include tetraglyme, pentane, cyclohexane, heptane, benzene, xylene, toluene, diethyl ether, tetrahydrofuran, butyraldehyde, and benzonitrile. Organic solvents may also contain dissolved water up to saturation limits. Exemplary preferred solvents include ketones (e.g., acetone and methyl ethyl ketone), esters (e.g., ethyl acetate, di-2-ethylhexyl phthalate, 2,2,4-trimethyl-1,3-pentanediol monoisobutyrate). , hydrocarbons (eg, toluene), nitrohydrocarbons (eg, nitrobenzene), ethers (eg, tetrahydrofuran (THF)), and sulfolane. In rhodium-catalyzed hydroformylation processes, the aldehyde product desired to be produced and/or For example, it may be desirable to employ as the main solvent an aldehyde compound corresponding to the higher boiling aldehyde liquid condensation by-product that can be produced in situ during the hydroformylation process. The primary solvent usually ends up containing both the aldehyde product and the higher boiling aldehyde liquid condensation by-product (the "heavies") due to the nature of the continuous process. The amount of solvent is not critical, as long as it is sufficient to provide the desired amount of transition metal concentration in the reaction medium. Typically, the amount of solvent ranges from about 5 weight percent to about 95 weight percent, based on the total weight of the reaction fluids. Mixtures of solvents may also be used.

Embodiments of the present invention are applicable to improving the conventional continuous mixed gas/liquid phase CSTR rhodium-phosphorus complex catalyzed hydroformylation process for producing aldehydes, conducted in the presence of free organophosphorus ligands. be. Such hydroformylation processes (also referred to as "oxo" processes) and their conditions are described, for example, in the continuous liquid recirculation process of U.S. Pat. No. 4,148,830, in U.S. Pat. , 668,651, are well known in the art. Also included are methods such as those described in US Pat. No. 5,932,772 and US Pat. No. 5,952,530. Such hydroformylation processes generally involve oxidizing olefinic compounds with hydrogen and monoxide in a liquid reaction medium containing a soluble rhodium-organophosphorus complex catalyst, a free organophosphorus ligand, and a higher boiling aldehyde condensation byproduct. It involves the formation of aldehydes by reaction with carbon gas. Generally, rhodium metal concentrations ranging from about 10 ppm to about 1000 ppm by weight, calculated as free metal, are sufficient for most hydroformylation processes. Some methods use about 10 to 700 ppm by weight of rhodium, often 25 to 500 ppm by weight, calculated as free metal.

Thus, as in the case of rhodium-organophosphorus complex catalysts, any conventional organophosphorus ligand can be used as the free ligand, and such ligands, and methods for their preparation, are well within the skill of the art. Well known in the field. A wide variety of organophosphorus ligands can be used in the present invention. Examples include, but are not limited to, phosphine, phosphite, phosphinophosphite, bisphosphite, phosphonite, bisphosphonite, phosphinite, phosphoramidite, phosphinophosphoramidite, bisphosphoramidite, fluorophosphite, etc. Not done. The ligand may contain chelate structures and/or multiple P(III) moieties such as polyphosphites, polyphosphoramidites, and mixed P(III) moieties such as phosphite-phosphoramidites, flurophosphite-phosphites, etc. III) may also be included. Of course, mixtures of such ligands can also be used if desired. Thus, the hydroformylation process of the invention may be carried out with any excess amount of free phosphorus ligand, for example at least 0.01 mole of free phosphorus ligand per mole of rhodium metal present in the reaction medium. can. Generally, the amount of free organophosphorus ligand used depends solely on the desired aldehyde product and the olefin and complex catalyst used. Therefore, for most purposes, the amount of free phosphorus ligands present in the reaction medium ranges from about 0.01 to about 300 or more per mole of rhodium present (measured as free metal). Are suitable. For example, large amounts of free triarylphosphine coordination, such as more than 50 moles, and sometimes more than 100 moles of free ligand, per mole of rhodium are generally required to achieve satisfactory catalyst activity and/or catalyst stabilization. While other phosphorus ligands such as alkylarylphosphine and cycloalkylarylphosphine have been used, the amount of free ligand present in the reaction medium is greater than 1 mole of rhodium present. As little as 1 to 100 moles per molecule, and in some cases even 15 to 60 moles, can be used to provide acceptable catalyst stability and reactivity without unduly interfering with the conversion of a particular olefin to an aldehyde. It can be helpful. In addition, other phosphorus ligands (e.g., phosphines, sulfonated phosphines, phosphites, diorganophosphites, bisphosphites, phosphoramidites, phosphonites, fluorophosphites) can be added to the free coordination present in the reaction medium. Amounts of as little as 0.01 to 100 moles and, in some cases, 0.01 to 4 moles per mole of rhodium present, can be tolerated without interfering with the conversion of certain olefins to aldehydes. May help provide catalyst stability and reactivity.

More specifically, exemplary rhodium-phosphorous complex catalysts and exemplary free phosphorus ligands include, for example, U.S. Pat. , 247,486, 4,283,562, 4,400,548, 4,482,749, European Patent Application Publication No. 96,986, 96,987 and No. 96,988 (all published on December 28, 1983) and International Publication No. 80/01690 (published on August 21, 1980). Among the more preferred ligands and complex catalysts that may be mentioned are, for example, the triphenylphosphine arrangements of U.S. Pat. alkylphenylphosphine and cycloalkylphenylphosphine ligands and rhodium-alkylphenylphosphine and rhodium-cycloalkylphenylphosphine complex catalysts, U.S. Pat. No. 4,283,562; There are diorganophosphite ligands and rhodium-diorganophosphite complex catalysts of Patent No. 4,599,206 and US Pat. No. 4,668,651.

As further discussed above, hydroformylation reactions are typically carried out in the presence of higher boiling aldehyde condensation byproducts. The nature of the continuous hydroformylation reaction that can be used herein is such that during the hydroformylation process, as more fully described, for example, in U.S. Pat. No. 4,148,830 and U.S. Pat. Such higher boiling aldehyde byproducts (eg, dimers, trimers, and tetramers) are generated in situ. Such aldehyde byproducts provide excellent carriers for liquid catalyst recycling processes. For example, the hydroformylation reaction can first be carried out in the absence or presence of a small amount of a higher boiling aldehyde condensation by-product as a solvent for the rhodium complex catalyst, or the reaction can be carried out in the absence or presence of a total liquid reaction. It can be carried out in the presence of 70 weight percent or more, even 90 weight percent or more, based on the total medium, of such condensation byproducts. Generally, a weight ratio of aldehyde to higher boiling aldehyde condensation byproduct within the range of about 0.5:1 to about 20:1 is sufficient for most purposes. Additionally, it is to be understood that small amounts of other conventional organic co-solvents may be present, if desired.

As noted above, hydroformylation reaction conditions can vary over a wide range, but generally the combined gas pressure of hydrogen, carbon monoxide, and olefinically unsaturated starting compound is less than about 3100 kilopascals (kPa), and more preferably More preferably, the method is carried out at a pressure of less than about 2415 kPa. The minimum total pressure of the reactants is not particularly critical and is primarily limited only by the amounts of reactants necessary to obtain the desired reaction rate. More specifically, the carbon monoxide partial pressure of the hydroformylation reaction method of the present invention can be about 1 to 830 kPa, optionally about 20 to 620 kPa, while the hydrogen partial pressure can be about 30 to 1100 kPa, optionally Depending on the pressure, it can be about 65-700 kPa. Generally, the H ₂ :CO molar ratio of gaseous hydrogen to carbon monoxide ranges from about 1:10 to 100:1 or more, and even in some cases from about 1:1.4 to about 50:1. good.

Additionally, as noted above, the hydroformylation reaction process of the present invention can be carried out at a reaction temperature of about 50°C to about 145°C. However, in general, hydroformylation reactions at reaction temperatures of about 60°C to about 120°C, or about 65°C to about 115°C are typical.

Of course, the particular manner in which the hydroformylation reaction is carried out and the particular hydroformylation reaction conditions used are not strictly critical to the invention and may be widely varied and adjusted to suit individual needs to achieve specific desires. It should be understood that aldehyde products can be produced.

Since the internal cooling coil alone often lacks sufficient heat removal capacity (limited heat transfer area per coil volume), the external cooling loop (external heat exchanger (cooler) Pump circulation) is typically used for highly exothermic hydroformylation reactions such as low carbon olefins (C2 to C5). In addition, internal cooling coils displace the internal reactor volume, increasing reactor size for a given production rate. However, in some embodiments, at least one internal cooling coil is placed inside the reactor, typically in the first reaction zone. In some embodiments, such an internal cooling coil can be added to the external cooling loop. In a preferred embodiment, the liquid process fluid used (separately or with a modification of the high-shear microbubble generator) to generate the jet (preferably microbubbles) before being reintroduced into the same reaction zone. (before the generator) passes through a heat exchanger. The flow of the cooled process fluid can be varied to provide optimal temperature control of the reaction zone, for example as taught in US Pat. No. 9,670,122 (particularly FIG. 3).

Preferred examples of olefins that can be used as reactants in the present invention include ethylene, propylene, butene, 1-hexene, 1-octene, 1-nonene, 1-decene, 1-undecene, 1-tridecene, 1- Tetradecene, 1-pentadecene, 1-hexadecene, 1-heptadecene, 1-octadecene, 1-nonadecene, 1-eicosene, 2-butene, 2-methylpropene, 2-pentene, 2-hexene, 2-heptene, 2-ethyl Hexene, 2-octene, styrene, 3-phenyl-1-propene, 1,4-hexadiene, 1,7-octadiene, 3-cyclohexyl-1-butene, allyl acetate, allyl butyrate, methyl methacrylate, vinyl methyl ether, Examples include vinyl ethyl ether, allyl ethyl ether, n-propyl-7-octenoate, 3-butenenitrile, 5-hexenamide, 4-methylstyrene, and 4-isopropylstyrene. Mixtures of isomers (eg butene raffinate) can also be used. The resulting aldehyde product can be hydrogenated and converted to the corresponding alcohol, which can be used as a solvent or in the preparation of plasticizers, or aldol condensation to higher aldehydes, oxidation to the corresponding acid, or the corresponding Other subsequent reactions such as esterification to produce acetate, propionate, or acrylate esters may be performed.

The olefin starting material can be reacted by any convenient technique, either as a gas (optionally with the incoming synthesis gas feed), as a liquid in the reactor, or as part of a recycle loop before entering the reactor. introduced into the vessel. One particularly useful method uses a separate olefin sparger (described below) next to or below the jet or optional syngas sparger to introduce the olefin feed and the syngas feed in close proximity to each other. That's true.

To help explain the operation of some embodiments of the hydroformylation reaction method of the present invention, reference is now made to FIG. FIG. 1 shows a non-limiting example of a cylindrical reactor 1 that can be used to carry out a hydroformylation reaction method according to an embodiment of the invention. Reactor 1 contains a reaction fluid that is a mixture of olefin, hydrogen, carbon monoxide, homogeneous catalyst, aldehyde product, solvent, and other components. The reactor has three reaction zones 1A, 1B, 1C. A portion of the reaction fluid is removed from the first reaction zone 1A through the outlet 3 at the bottom of the reactor. At least a portion of the removed reaction fluid passes through two shear mixers 4 and is then mixed with fresh or recycled syngas (with or without inerts), as shown in FIG. 3a or 3b. is introduced to generate bubbles within the portion of the removed reaction fluid. The removed reaction fluid is returned to the first reaction zone 1A via the two nozzles 5. Nozzles and their orientation are discussed further below. The removed reaction fluid is returned to the first reaction zone 1A through a nozzle 5 that forms one or more liquid jets of returned reaction fluid that impart momentum and vapor to the majority of the reactor fluid. Gives a mixture. Shear mixing devices include those described in US Pat. No. 5,845,993, which is incorporated herein by reference.

With respect to the reaction fluid removed from the bottom of reactor 1 via stream 2, crude product and catalyst mixture can be removed from stream 2 via a product-catalyst separation zone (not shown). This stream 2 may also be passed through a heat removal process such that the returned process fluid is cooled to cool the reaction zone.

As used herein, the terms "shear mixing device," "high shear mixing device," "microbubble generator," and "high shear microbubble generator" are used interchangeably and include Refers to equipment that can generate bubbles with an average size of 3,000 microns or less. Important features and advantages of shear mixing devices that can be used in embodiments of the present invention are that the entire shear mixing device (e.g., does not contain moving parts or require mechanical seals, does not require maintenance, It is constructed from static piping components (with no potential for leakage/failure points), thus increasing inherent safety, mechanical reliability, low emissions to the environment, and plant uptime. An example of a shear mixing device that can be used in embodiments of the present invention is described in US Pat. No. 5,845,993, which is incorporated herein by reference. Generally, shear mixing devices include pressurized gas conduits or chambers in contact with single (or multiple) turbulent liquid streams separated by perforated surfaces. Gas enters the liquid stream through the perforations, driven by the shear stress generated by the liquid stream. Two typical embodiments of such shear mixing devices are shown in FIG. For example, in one embodiment (FIG. 3a), the shear mixing device has an internal channel carrying the liquid flow (L). It is fitted with an outer concentric jacket connected to the pressurized gas inlet (G). A portion of the inner channel enclosed by the outer jacket is perforated by multiple perforations. These perforations are where the gas (G) from the outer jacket enters the liquid (L) stream in the inner channel in the form of a gas-in-liquid dispersion composed of small air bubbles. In the present invention, the liquid (L) is at least a portion of the removed reaction fluid that is returned to the first reaction zone, and the gas (G) is synthesis gas.

In some embodiments, a portion of the syngas is introduced via conventional sparging (e.g., as disclosed in WO 2018/236823) in addition to the syngas introduced via a shear mixing device. may be introduced into the first reaction zone. In other embodiments, the only source of syngas provided to the first reaction zone is through a shear mixing device.

In embodiments of the invention, the bubbles are generated by a shear mixing device, so the mixing energy introduced into the first reaction zone without using conventional sparging is different from WO 2018/236823 . The momentum generated by the flow through the shear mixing device is required to evenly distribute the bubbles throughout the reaction fluid starting at the nozzle outlet. The majority of the momentum from the jet exiting the nozzle does not need to reach the bottom of the first reaction zone in some embodiments where conventional spargering is used, but instead drives air bubbles throughout the first reaction zone. It only needs to be distributed. In order to achieve suitable mixing and gas distribution, there are several reactor and nozzle design considerations that need to be addressed, as discussed further below.

As shown in the mixing energy density formula below, the inventors found that excellent results are achieved when the mixing energy density (power delivered per unit volume) provided by the jet exceeds 500 W/ ^m3 . I discovered that. In the absence of such mixing energy, there is less (or no) turbulence in the reacting fluid, so the gas bubbles increase in size and rise rapidly to the gas-liquid interface due to increased buoyancy and detachment from the liquid. As a result, gas stagnation within the reactor is reduced. Creating and maintaining small gas bubbles is important for producing a uniform reaction fluid that provides better gas-liquid mixing, gas stagnation, and more reproducible reactor performance. The smaller bubbles provide maximum gas stagnation and maximize the mass transfer area between the bubbles and the liquid dissolving the syngas (optimized gas volume/surface ratio). Conversely, very small air bubbles may become trapped within the liquid streamline near the outlet nozzle (e.g. to an external recirculation pump/heat exchanger or at the reactor product outlet); An important feature of the present invention is the ability to consistently produce bubbles in the appropriate size range, as downstream equipment may be adversely affected.

Referring again to Figure 1, reaction fluid is removed from the bottom of the reactor 1 via outlet 3 and via two or more nozzles 5, optionally terminating in a flow divider plate or restriction nozzle (described below). returned to the reactor. In some embodiments, two or more nozzles 5 may be oriented in symmetrical pairs, symmetrical triads, or other symmetrical arrangements.

The nozzle 5 may be oriented to direct the liquid jet downwardly or upwardly or both. In some embodiments, the nozzle may be oriented such that the liquid jet is not directed toward the central longitudinal axis of the reactor 1 (eg, not directed toward the central axis of the reactor). Preferably, the liquid jet is not oriented strictly horizontally or strictly vertically or directly towards the vertical axis or center of the reactor. Nozzle orientation is further discussed below in connection with FIG. 2.

In some embodiments, multiple sets of symmetrical nozzles may be arranged at different nozzle orientations (radial positions) and/or at different heights within the reactor 1. In some embodiments, various liquid feeds (e.g., liquid olefin feed, liquid catalyst stream from an upstream reactor, liquid catalyst stream from a product-catalyst separation zone, etc.) are routed through shear mixing device 4. can be provided to reactor 1. In some embodiments, one or more of such feeds may be combined with returned removed reaction fluid and provided to reactor 1 via at least one shear mixing device. If the liquid feed is from an upstream reactor, some syngas may be present, but this represents a small amount of syngas compared to the syngas introduced by shear mixer 4. In the embodiment shown in FIG. 1, fresh liquid olefin feed 6 is combined with returned reaction fluid 7 and provided to reactor 1 via a shear mixing device.

The returned removed reaction fluid exits each nozzle 5 as a jet. As used herein, the terms "jet," "directed jet," and "directed flow" are used interchangeably to indicate that syngas is produced by one or more shear mixing devices rather than by sparging. As described in WO 2018/23623, except as provided. The jet is the output of one or more shear mixing devices or a separate stream specifically designed to mix the first reaction zone (separately or in conjunction with the shear mixing device). Good too.

The jets provide a downward counterflow to offset the natural buoyancy of the bubbles and maintain entrainment of the bubbles in the liquid circulating throughout the backmixing reactor, resulting in the increase of syngas bubbles throughout the backmixing liquid phase. resulting in a more even distribution. As the syngas dissolves and reacts, the gas bubbles contract, which further helps maintain distribution within the back-mixed liquid phase and promote good gas mass transfer into the liquid phase. This homogeneously mixed liquid reaction fluid moves across a permeable physical barrier, such as a perforated partition (discussed below), into the unstirred reaction zone, thus providing external mixing energy in some embodiments. React in a controlled manner without need.

The returned jet of reactant fluid provides a mixing energy density to the reactant fluid to properly mix the reactants in the reactant fluid and promote the reaction.

In some embodiments, the jets provide sufficient mixing energy density such that an agitator or other mechanical source of mechanical mixing energy is not required. The jet has the following formula:

(where V is the volume of reaction fluid in the first reaction zone (m ³ ) and N is the volume of reaction fluid in the first reaction zone (m 3 ), and N is the is the total number of jets returned to the first reaction zone, as identified in , and ρ _i is the average density (kg /m ³ ), Q _i is the volumetric flow rate (m 3 /s) of the reaction fluid returned to the first reaction zone through the i-th jet, and A _i is the volumetric flow rate (m ³ /s) of the reaction fluid flowing through the i-th of the nozzle at the location where the reactant fluid exits the nozzle and enters the ^first reaction zone. For clarity, V (volume of reaction fluid in the first reaction zone (m ³ )) is the gas-filled liquid level when the method is carried out (as opposed to the evacuated liquid volume). refers to This volume (V) can be determined by known methods such as sonar level indicators or take-off nozzles. Similarly, ρ _i can be easily calculated by the relative flow rates of the reaction fluid and synthesis gas fed to the shear mixing device. Average density (ρ) at the nozzle port of the reaction fluid returned to the first reaction zone through the i-th jet (kg/m ³ ), Reaction fluid returned to the first reaction zone through the i-th jet The volumetric flow rate (Q _i ) (m ³ /s) of and the cross-sectional area (A _i ) (m ² ) of the i-th nozzle through which the reaction fluid flows are determined by techniques known to those skilled in the art based on the teachings herein. can be measured or determined using The jet is believed to provide adequate mixing in the first reaction zone by providing a mixing density energy (as defined by the above formula) of 500 watts/m ³ or more. In other words, in some embodiments, the jets can mix well without the need for conventional mechanical agitators.

The flow rate of the reaction fluid through the shear mixing device may also be important to ensure that adequate mixing energy is provided to the first reaction zone. Accordingly, in some embodiments, the flow rate of the reactant fluid through the shear mixing device is:
q _SM >525 (μ _o /ρ _o ) P _SM
(where q _SM is the flow rate of the reaction fluid entering the shear mixing device (m ³ /s), ρ _o is the density of the reaction fluid before entering the shear mixing device (kg/m ³ ), μ _o is the viscosity of the reaction fluid before entering the shear mixing device (Pa·s) and P _SM is the minimum wetted perimeter of the cross section of the liquid flow inside the shear mixing device. The flow rate of the reaction fluid entering the shear mixer (q _SM ) (m ³ /s), the density of the reaction fluid before entering the shear mixer (ρ _o ) (kg/m ³ ), and the density of the reaction fluid before entering the shear mixer (kg/m 3 ); The viscosity (μ _o ) (Pa·s) of the reaction fluid can be measured using techniques known to those skilled in the art based on the teachings herein. The minimum wetted perimeter (P _SM ) of the cross section of the liquid flow inside the shear mixing device can be determined as follows. For conventional tubes transporting reactant fluids through a shear mixing device, P _SM is π multiplied by the inner diameter of the tube (P _SM = ID _tube ). Optionally, an inner tube may also be present, and the reaction fluid may flow in an annular region between the outer wall of the inner tube and the inner wall of the outer tube. In this case, P _SM is π multiplied by the sum of the outer diameter of the inner tube and the inner diameter of the outer tube (P _SM =π(OD _{inner tube} + ID _{outer tube} )).

In one embodiment, all of the jets are from a shear mixing device. In other embodiments, some jets are solely for imparting mixing energy density and other jets are from one or more shear mixing devices. In another embodiment, a multi-zone reactor has shear mixer jet loops in multiple reaction zones within the reactor, each jet loop recycling fluid removed from the same zone that was removed. Circulate. In another embodiment, the multi-zone reactor can be configured to remove reaction fluid from a first reaction zone and return the reaction fluid as a jet to a second reaction zone via a shear mixing device. . In a further embodiment, all zones within the reactor body have jets from a high shear mixing device. In a preferred embodiment, the second reaction zone is not a backmix reactor, but a bubble column reactor, a plug flow reactor, a piston flow reactor, a gas or bubble lift (tubular) reactor, a packed bed reactor, or selected from venturi type reactors. Examples of non-backmixing reactors include U.S. Pat. No. 5,367,106, U.S. Pat. No. 5,105,018, U.S. Pat. .

The position and orientation of the nozzle within the reactor is important, especially when more than one nozzle is provided. For example, arranging two nozzles so that the jets exiting the nozzles are directly directed toward each other should generally be avoided. FIG. 2 is a schematic diagram of a side view and two top views of a cylindrical reactor 100 to illustrate the position and orientation of the nozzle 105 according to some embodiments of the invention. FIG. 2 also shows a donut baffle 110 (discussed further below) located below the nozzle 105 within the reactor 100. Alpha (α) is the angle of the nozzle with respect to the horizontal plane. In some embodiments, when the horizontal angle is 0°, α can range between 75° (upward angle) and −75° (downward angle). In some embodiments, when the horizontal angle is 0°, α can range between 45° (upward angle) and −60° (downward angle). Beta (β) is the angle at which the nozzle is oriented to the left or right with respect to the centerline. In some embodiments, when the orientation of the nozzle pointing straight across the reactor is 0°, β is approximately 5° to 90° (nozzle oriented clockwise when viewed from the top of the reactor). or -5° to -90° (nozzle oriented counterclockwise when viewed from the top of the reactor). β should be between -5° and 5° if phi (φ) is greater than 0°. Phi (φ) is the distance that the nozzle is offset from the centerline of the reactor when viewed from above. φ should be no more than 50% of the cross-sectional diameter of the reactor. In some embodiments in which at least two nozzles return the removed reaction fluid to the reactor, each nozzle is oriented such that the angle of the nozzle with respect to the horizontal plane (alpha (α)) is between +75° and -75°. , alpha (α), the angle of the nozzle with respect to the vertical plane passing through the center of the reactor (beta (β)), and the distance from the vertical plane passing through the center of the reactor (phi (φ)) when beta is zero. are all non-zero.

In some embodiments, the flow of the returned reaction fluid is not a single line, but the majority of the reaction fluid returned to the reactor in a single nozzle is within a relatively narrow range of α and β values. I hope you understand that. For the purposes of this application, when the terms "vertical" and "horizontal" are used in connection with the flow of the returned reaction fluid in a fluid diverter, those terms use angles α and β, respectively. can be understood as such. That is, a "vertical flow" or "vertical jet" is directed upward and/or downward with non-zero α and essentially zero β. A "horizontal flow" or "horizontal jet" is oriented to the left and/or to the right with essentially zero α and non-zero β. The term "directed flow" generally refers to a flow in which both α and β are non-zero. "Directed flow" may include flow from the shear mixing device or other flow that returns to but does not pass through the shear mixing device.

Still referring to FIG. 2, delta (δ) is the distance that the nozzle projects from the reactor wall into the reactor. In some embodiments, δ is less than 50% of the diameter of the reactor. In some embodiments, δ is 50% or less of the radius of the cylindrical reactor. In some embodiments, δ is at least 10% of the radius of the cylindrical reactor. In some embodiments, δ is 10% to 45% of the radius of the cylindrical reactor. In some embodiments, the end of the flow divider can be generally flush with the reactor wall such that δ is about 0% of the radius of the cylindrical reactor.

In some embodiments, additional sets of nozzles can be provided at the same or different heights, as shown in FIG. 2, or at different angles (α and/or β). FIG. 4 illustrates different positions of nozzle 205 within reactor 200, one for jets (not labeled but represented by arrows exiting ports of nozzle 205), according to some embodiments of the invention. FIG. 3 is a series of views showing different positions of the two or more donut baffles 210 and the angle of the jet exiting the nozzle 205. A shear mixing device 215 is also shown, but not numbered in the figures of FIG.

Phi (ψ) is the distance at which the tip of the nozzle is located (as a percentage of the reactant fluid filling height). As used herein, "reactant fluid fill height" refers to the height of liquid within the reactor from the bottom of the reactor. In embodiments where the reactor has a headspace in the bottom portion, as shown in Figure 2, all heights referred to as measured from the bottom of the reactor are across the reactor directly above the headspace. Measured from tangent 102. As further shown in Figure 2, if the reactor has a flat bottom, all heights referred to as being measured from the bottom of the reactor are measured from the physical bottom. ψ is less than 100% of the reaction fluid fill height. In some embodiments, ψ is less than 90% of the reactant fluid fill height. In some embodiments, ψ is 75% to 80% of the reactant fluid fill height.

Each shear mixing device is designed to introduce synthesis gas bubbles into the removed reaction fluid. Without being bound by theory, the high liquid velocity and thorough mixing with small initial bubble size provided by embodiments of the present invention minimize synthesis gas bubble coalescence and reduce bubble size due to shear. Facilitates shrinkage and uniformity of gas-liquid distribution and temperature throughout the reaction zone. The movement of small syngas bubbles due to their own natural buoyancy is offset by the viscosity of the liquid and the turbulence of the liquid mass. Similarly, if the unstirred reaction zone is above the first reaction zone, the natural buoyancy force rising up to and crossing a permeable physical barrier, such as a grid or perforated plate, separating the two zones is due to the viscosity of the liquid. and is offset by the turbulence of the liquid mass. Excessively large bubbles will rise too quickly, resulting in less gas stagnation and uneven distribution. In some embodiments, the average size of the bubbles produced by the shear mixing device can be between 10 nanometers and 3,000 microns. In some embodiments, the bubbles produced by the shear mixing device average between 3 microns and 3,000 microns. In some embodiments, the average bubble size produced by the shear mixing device is between 30 microns and 3,000 microns. In some embodiments, the average size of the bubbles produced by the shear mixing device is between 100 microns and 800 microns.

The manner in which the reaction fluids are returned affects the effectiveness of the mixing energy provided. In some embodiments, the reactant fluid is returned using piping that has one or more flow divider plates installed at the end of the piping and inserted through and attached to the recirculation return nozzle of the reactor. You can. In some embodiments, the reactant fluid is returned using a nozzle or flow orifice located at the end of the piping and inserted through the recirculation return nozzle of the reactor and into it, as described further below. It is attached. In each case, the velocity of the resulting liquid jet is a function of the flow area of the nozzle or orifice, or the flow area created between the diverter plate and the inner wall of the piping, and the mass flow rate and density of the returned reaction fluid. It is. The combination of flow area and flow rate creates a jet of reactant fluid within the reactor, which imparts momentum and causes gas-liquid and liquid-liquid mixing of the bulk fluid within the reactor. Furthermore, the returned reaction fluid is split and directed in multiple directions.

The term "flow divider" herein encompasses both nozzles and flow divider plates located within the reactor recycle return pipe. In either case, the flow divider directs the flow of the returned reaction fluid. As described further below, in some embodiments, the flow divider directs the flow of the returned reaction fluid horizontally. In some embodiments, the flow divider directs the flow of the returned reaction fluid vertically. In some embodiments, the flow divider directs the flow of the returned reaction fluid both horizontally and vertically. Flow dividers comprising a flow divider plate located at the end of a tube are described in further detail in WO 2018/236823, which is incorporated herein by reference.

In some embodiments, horizontal donut baffles above or below the nozzle are used to reduce flow or channeling effects in the reactor from the jet. A donut baffle is a flat ring plate fixed to the reactor wall with a central opening that functions to divert the channeling flow along the reactor wall. Non-limiting examples of such horizontal donut baffle arrangements are shown in FIG. 1 (reference number 14), FIG. 2 (reference number 110) and FIG. 4 (reference number 210). As shown in FIG. 2, donut baffle 110 extends a distance (γ) from the reactor wall. In some embodiments, the donut baffle extends from the reactor wall a distance (γ) that is between 5% and 25% of the reactor diameter. In some embodiments, the vertical position of the donut baffle within the reactor may also be important. As shown in FIG. 2, the donut baffle 110 is placed at a particular height (λ) from the bottom of the reactor. In some embodiments, the donut baffle may be placed at a height from the bottom of the reactor (λ) that is 90% or less of the reaction fluid fill height. Other techniques may be used to minimize the potential for flow or channeling effects from the reactant fluid entering the reactor as a jet (e.g., donut baffle 14 in FIG. 1 and donut baffle 210 in FIG. 4). (Please refer to the location).

In some embodiments, such as the embodiment shown in FIG. 1, the reaction fluid passes from the first reaction zone through a product outlet nozzle 3 at the bottom of the reactor. In such embodiments, the reactor may include gas entrainment prevention means 8 located in the bottom volume of the reactor. Such means may be in the form of an entrained separator, a conical condenser, one or more perforated plates, or a packed bed. A packed bed 8 is shown in FIG. Such gas entrainment means 8 are particularly desirable when the jet is downwardly inclined, but may be unnecessary if the recirculation pump can tolerate small entrained gas bubbles.

In some embodiments, when a single reactor includes multiple reaction zones, perforated dividers can be placed between the reaction zones. For example, as shown in FIG. 1, the reaction fluid from the first reaction zone 1A passes upward through the perforated partition plate 10 and enters the second reaction zone 1B. The perforated partition plate 10 may help ensure that the reaction fluid rising into the second reaction zone 1B is homogeneous and contains a significant amount of synthesis gas for continued reaction. This is particularly desirable for bubble reactors, plug flow reactors, and packed column reactors in that the reactants are highly homogeneous and diffusion is not restricted.

In the embodiment shown in FIG. 1, a second perforated partition plate 12 separates the second reaction zone 1B from the third reaction zone 1C.

To be effective, the perforations in the perforated partition should be evenly distributed to evenly distribute the rising fluid across the cross section of the reactor. In plug flow reactors or packed bed column reactors, the perforations should direct the flow so that each tube or column gets the same fluid flow. Perforated divider or tray designs are well known in the art. Typical perforated dividers/trays should have a porosity of 15-40% (preferably 20-30%), with perforations evenly distributed over the surface. The perforations may be uniform or have different diameters with equivalent hole diameters typically ranging from 1/8 inch to 2 inches. The holes may be circular, square, slotted, or other shapes, and may have additional features (e.g., countersinks, raised holes, etc.), but do not allow large amounts of gas to flow under the perforated partition. It must not be something that accumulates. In some embodiments, wire mesh or similar rigidly supported materials may be used as an alternative to perforated dividers.

Vertical baffles may be attached to the interior walls of the first reaction zone to shear and lift radial streamlines from the walls of the vessel to provide additional mixing and minimize rotational flow.

Returning to FIG. 1, in the final reaction zone 1C there is a reactor outlet 9 for conveying the reaction fluid to the next reactor or product-catalyst separation zone (not shown).

Additionally, in embodiments where multiple reactors are arranged in series, the optional gas purge stream 14 from reactor 1 may be discharged, spread out, and sent to a plant fuel gas header or another reactor. Analysis of this purge stream 14 can provide a convenient means of measuring the CO partial pressure in the upper reaction zone for reaction control.

Although not shown in Figure 1, the system may include other standard components that can be readily recognized and implemented by those skilled in the art, such as pumps, heat exchangers, cooling coils, valves, level sensors, temperature sensors, and pressure sensors. It also includes parts for mechanical equipment.

In some embodiments, the removed reaction fluid returned to the first reaction zone through one or more nozzles can provide at least 50% of the total mixing energy density to the first reaction zone. In some embodiments, the removed reaction fluid returned to the first reaction zone through one or more nozzles can provide at least 85% of the total mixing energy density to the first reaction zone. In some embodiments, the returned reaction fluid can provide substantially all or 100% of the total mixing energy density to the first reaction zone. The total mixing energy density includes the mixing energy provided by the operating agitator (if present), jets of returned reactant fluids, or other mixing energy density sources, but does not include the mixing energy provided by the operating agitator (if present), jets of returned reactant fluids, or other sources of mixing energy density, but It should be understood that any of the insignificant mixing energy densities provided by introducing the substances into the reactor are not included. For example, there is some negligible mixing energy density provided by the hydraulic flow of liquid reaction fluid from a first reaction zone through a subsequent reaction zone (e.g., through a divider, if present). However, this is not included. In embodiments where the liquid jet produced by the returned reaction fluid provides substantially all or 100% of the mixing energy density, the reactor does not include an agitator or includes an inactive agitator.

If an agitator is present and operating, the mixing energy density contribution from the agitator is given by the following equation:
P=N _pg ×ρN ³ D ⁵
(In the formula, N _pg is the gassed power number of the impeller, ρ is the density of the reaction fluid, N is the rotational speed of the stirrer (rev/s), and D is the diameter of the impeller.) can do.

Surprisingly, by using the shear mixing device described herein, the hydroformylation reactor can be heated without operating an agitator while providing the same level of gas-liquid and liquid-liquid mixing of the reaction fluids. It turns out that it is possible to make the operation possible. Increasing the flow rate of the returned reaction fluid can allow stable operation without activating the agitator and facilitate better gas dispersion into the liquefied reaction fluid. Some embodiments of the invention can be used in a reactor without the use of an agitator when there are agitator motors, agitator seals, agitator shafts/impellers, stationary bearings or similar agitator-related issues. Providing proper mixing can allow continuous operation of a reactor with an agitator until the unit can be shut down and repairs can be made, thus eliminating unplanned production losses. It has the advantage of avoiding In other words, for retrofits, some embodiments of the present invention may allow existing agitators to be repaired without operation and/or while the reactor continues to operate. For new reactors, some embodiments of the present invention can advantageously eliminate the cost of agitators and the need for agitator seals and stable bearings that require maintenance/replacement, and can eliminate seal leaks. .

Some embodiments of the invention will now be described in detail in the following examples.

In this example, the performance of three designs was evaluated using computational fluid dynamics ("CFD") tools. Comparative Example A is representative of the prior art using a mechanical stirrer. Examples 1 and 2 of the invention represent embodiments of the invention that utilize a shear mixing device without a mechanical stirrer. The objective was to determine the comparability and performance criteria of the agitator-free designs of the present invention (Inventive Example 1 and Inventive Example 2) with respect to the conventional mechanically agitated design (Comparative Example A). / or to show improvement. In this example, CFD was used to evaluate (a) mixing efficiency (i.e., mixing time), (b) gas dispersion (i.e., uniformity of gas volume fraction and overall gas stagnation), and (c) desorption. (i.e., the volume percent of gas in the bottom recirculation line); and (d) mass transfer (i.e., the average value of the mass transfer coefficient (KLa) in the first reaction zone).

Importance of Syngas Dispersion It is important to have highly and well-dispersed syngas within the reactor for several reasons. Only synthesis gas that is dispersed and dissolved in the reaction fluid can react, so the synthesis gas introduced into the reactor is quickly dispersed and dissolved in the reaction fluid instead of rising to the gas-liquid interface as bubbles, and It enters the vapor space and becomes unusable for the reaction. Additionally, the volume within the reactor in which syngas is not dispersed or dissolved does not contribute to the reaction or reactor productivity because it is starved of reactants. Therefore, highly dispersed (high gas stagnation or gas fraction) and uniform gas mixing is the most desirable outcome.

Method for Evaluating Effectiveness in CFD To evaluate the mixing properties of the present invention, it is convenient to examine the gas distribution from CFD modeling to determine both the uniformity of the gas distribution and the range of gas loading. Commercial experience with well-stirred CSTR reactors is that gas loading values range from 5 to 12%. CFD modeling programs can be used to predict overall or average gas loading values across the reactor, which de-emphasizes the local effects of areas of high or low gas loading and short residence times. (e.g. piping inlet/outlet, nearby agitator impeller, etc.)

Mixing effectiveness: mixing time θ _mix
- For well-mixed systems, the mixing time θ _mix should typically be less than 10% to 20% of the average liquid residence time θ _res . (See: Paul, E. L., V. A. Atiemo-Obeng and S. M. Kresta, eds. 2004. Handbook of Industrial Mixing: Science and Practice. John Wiley & Sons, I nc.).
- In this CFD simulation, the mixing time θ _mix was evaluated for the first reaction zone in each example.
・A well-known tracer injection method was implemented. The simulation was first run without tracer. After reaching steady state, the tracer was continuously injected into the fresh feed inlet and its concentration was followed in the first reaction zone. At each simulation time step, the coefficient of variation (CoV) was evaluated as the volume standard deviation of the concentration relative to its volume average.
- Mixing time was defined as the flow time to reach 5% CoV.

Gas distribution: uniformity of gas volume fraction and overall gas stagnation To evaluate the mixing properties of the present invention, gas distribution was investigated from CFD modeling to determine both the uniformity of gas distribution and the degree of gas loading. It is convenient to specify.
- Commercial experience with well-stirred CSTR reactors is that gas loading values range from 5 to 12%.

Degassing: Volume % of gas in the bottom recirculation line
- The bottom outlet of the reaction vessel typically leads to a centrifugal pump used to recirculate fluids.
- Synthesis gas bubbles may be entrained at the outlet if they are small enough. If the entrainment is large enough, the presence of gas can damage the pump.
- For safe pump operation, it is essential to keep the gas volume fraction in the bottom recirculation line below 3% to 5%.
- CFD modeling tracked this volume fraction and reported this value for each case.

Mass transfer:
- The overall effectiveness of the reaction in a gas-liquid system, such as a hydroformylation system, is based on the rate at which the syngas components (CO and H2) are transferred to the liquid phase.
- The mass transfer rate of synthesis gas into the liquid phase in any reaction zone is directly proportional to the average value of the mass transfer coefficient (kLa) in that reaction zone.
- In this example, CFD modeling was used to directly obtain the average value of the mass transfer coefficient (KLa) in the first reaction zone. For this purpose, methods well described in the literature were used. Gimbun, Reilly and Nagy "Modeling of mass transfer in gas-liquid stirred tanks agitated by Rushton turbine and CD-6 impell er: A scale-up study,” Chemical engineering research and design 87 (2009) 437-451.
- In order to equalize the performance of two reactors carrying out the same gas-liquid reaction under the same operating and feeding conditions, the overall volume KLa values should be equivalent. Generally, higher KLa values are preferred.

Operating Conditions and Parameters The following operating conditions and parameters were used for each of the examples. The operating pressure is approximately 15 bar (absolute). At this pressure, the density of liquid propylene is approximately 775 kg/m ³ and the density of syngas is approximately 9.06 kg/m ³ . Table 1 also shows the feed flow rates of synthesis gas and liquid propylene for each example. The viscosity of liquid propylene is 3.8×10 ⁻⁴ Pa. s, the viscosity of the synthesis gas is 1.8×10 ⁻⁵ Pa.s. Let it be s. The gas-liquid surface tension between the syngas and liquid propylene is 18 dynes/cm (0.018 N/m), in line with typical values for similar organics.

Comparative example A
The original reactor was a mechanically stirred tank with a diameter of 5.5 meters and a height of 10 meters in the cylindrical section, capped at the top and bottom by two identical 2:1 semi-ellipsoidal heads. Ta. The tank volume was vertically divided into three reaction zones (numbered 1-3 from bottom to top) by two horizontal baffles. The baffle is an identical stainless steel plate with the same diameter as the tank and a single central orifice 0.7 meters in diameter. Additionally, the tank was fitted with four identical vertical baffles spaced 90° apart along the reactor wall.

Synthesis gas was introduced using two identical ring spargers located in the first reaction zone (0.2 m above the lower tangent) and the second reaction zone (0.2 m above the lower horizontal baffle). The agitator is a shaft fitted with three impellers: a standard gas distribution turbine in the bottom compartment and two hydrofoils in the second and third reaction zones. The stirrer operates at 89 rpm.

A degassing ring concentric with the reactor body was attached to the bottom dished head around the bottom recirculation nozzle.

Table 1 summarizes the reactor dimensions and flow rates.

Example 1 of the present invention
Reactor dimensions (diameter and L/D) are the same as in Comparative Example A. There was no stirrer; instead, a recirculating jet entering the first reaction zone was used to perform mixing and gas dispersion. In addition, the following other changes were made to Comparative Example A.
1. Liquid recirculation flow rate was increased by a factor of 16.
2. Gas introduction and bubble size: The sparger ring was removed. Here, syngas was introduced directly into the recycle stream using two shear mixers, each located just upstream of the recycle inlet nozzle. These shear mixers were designed to introduce gas into the reactor with an average bubble diameter of 300 microns. Details of the shear mixing device are described in the shear mixing device section at the end of the Examples section.
3. Horizontal Baffle: The horizontal baffle of the Comparative Example A design was replaced with a stainless steel perforated plate. These plates have a 20% open area to allow the two-phase (gas-liquid) reaction fluid to pass vertically upward from the bottom compartment to the middle compartment.
4. Nozzle: Removed wedge insert from recirculation inlet nozzle. Each recirculation inlet nozzle includes a bend at the end, thereby
a. The gas-liquid jet enters at a vertical angle (α) of 20 degrees (downward along the central axis of the reactor) and an azimuthal angle (β) of 30 degrees (counterclockwise about the central axis of the reactor);
b. reducing the nozzle diameter to a nominal size of 7 inches at the end using a standard conical reducer (12 inches x 7 inches);
c. The nozzle opening through which the two-phase jet entered the first reaction zone was located at a height (δ) 2.52 m above the lower tangent of the reactor and 0.45 m radially inward from the inner wall of the reaction vessel.
5. Internal degassing: Degassing ring was removed. A packed bed with a porosity of 36% and a total height of 1.375 m was installed in its place.
6. Internal: A donut baffle was added to the first reaction zone to prevent channeling of gas to the second reaction zone. A donut baffle with a width of 0.59 m was placed 2 m above the lower tangent.
7. Bottom recirculation nozzle: The nozzle size was increased from the original size of 16 inches to 22 inches to reduce the liquid velocity exiting the reactor.

Table 2 summarizes various dimensions and other parameters of Example 1 of the present invention.

As mentioned above, FIG. 2 defines various parameters for characterizing the orientation and position of the nozzle within the reactor. Table 3 provides the values of these parameters for Examples 1 and 2 of the present invention.

Example 2 of the present invention
Example 2 of the present invention is the same as Example 1 of the present invention except for the following changes.
1. Internal degassing: the packed bed was removed. In its place, a stack of three horizontal perforated plates was used. The opening area fractions are 30% (upper layer plate), 20% (middle layer plate), and 15% (lower layer plate). The gap between the plates is 0.25 m.

Results The results of the CFD modeling are shown in Tables 4A and 4B.

As shown in Table 4, Examples 1 and 2 of the present invention had comparable performance (e.g., mixing time, KLa, and volume % of the gas in the circulation line). Examples 1 and 2 of the invention also have significantly lower power consumption (see P/V).

5 and 6 show gas volume fraction color-coded diagrams and KLa color-coded diagrams for Comparative Example A and Examples 1 and 2 of the present invention. As shown in FIG. 5, embodiments of the invention have very uniform gas volume fractions.

Shear Mixing Device The shear mixing device used in Examples 1 and 2 of the present invention is of the type described in US Pat. No. 5,845,993. Each device consists of a pressurized gas conduit or chamber in contact with a single (or multiple) turbulent liquid stream separated by a perforated surface. Gas enters the liquid stream through the perforations, driven by the shear stress generated by the liquid stream.

For Examples 1 and 2 of the present invention, as shown in Figure 3a, the shear mixing device is composed of internal channels carrying liquid flow. This was fitted with an outer concentric jacket connected to a pressurized gas inlet. A portion of the inner channel enclosed by the outer jacket is perforated by multiple perforations. These perforations are where gas from the outer jacket enters the liquid flow within the inner channel in the form of a gas-in-liquid dispersion composed of small air bubbles. In Examples 1 and 2 of the invention, the liquid is the reaction fluid removed from the reactor and the gas is synthesis gas.

The shear mixing device was configured to provide an average cell size of 300 microns. The dimensions and flow rates in the shear mixing device to provide this average cell size are shown in Table 5.

Claims (16)

A hydroformylation reaction method, comprising:
(a) contacting an olefin, hydrogen, and carbon monoxide in the presence of a homogeneous catalyst in a reactor to provide a reaction fluid, the reactor comprising one or more reaction zones; ,
(b) removing a portion of the reaction fluid from the first reaction zone;
(c) passing at least a portion of the removed reaction fluid through a shear mixing device to generate gas bubbles in the portion of the removed reaction fluid, the hydrogen and hydrogen provided to the reactor at least a portion of the carbon oxide is introduced through the shear mixing device;
(d) returning the removed reaction fluid to the first reaction zone through one or more nozzles, the returned reaction fluid exiting each nozzle being a jet;
including;
The mixing energy density provided to the reactor by the jet is determined by the following equation:
(where V is the volume of the reaction fluid in the first reaction zone (m ³ ) and N is the volume of the reaction fluid in the first reaction zone (m 3 ), and N is is the total number of jets returned to the first reaction zone to be uniquely identified, and ρ _i is the average density ( kg/m ³ ), Q _i is the volumetric flow rate (m 3 /s) of the reaction fluid returned to the first reaction zone through the i-th jet, and A _i is the volumetric flow rate (m ³ /s) of the reaction fluid returned to the first reaction zone through the i-th jet. is the cross-sectional area (m ² ) of the i-th flowing nozzle at the location where the reaction fluid exits the nozzle and enters the first reaction zone)
satisfy,
Method. The flow rate of the reaction fluid through the shear mixing device is:
q _SM >525 (μ _o /ρ _o ) P _SM
(where q _SM is the flow rate (m ³ /s) of the reaction fluid entering the shear mixing device, and ρ _o is the density (kg/m ^{3 )} of the reaction fluid before entering the shear mixing device. ), μ _o is the viscosity (Pa·s) of the reaction fluid before entering the shear mixing device, and P _SM is the minimum wetted perimeter of the cross section of the liquid flow inside the shear mixing device. )
The method according to claim 1, wherein the method satisfies the following. At least two nozzles return the removed reaction fluid to the reactor, each nozzle oriented such that the angle (alpha) of the nozzle with respect to a horizontal plane is between +75° and −75°, and alpha, the 1 . The angle of the nozzle with respect to a vertical plane passing through the center of the reactor (beta) and the distance from the vertical plane passing through the center of the reactor (phi) when beta is zero are all non-zero. Or the method described in 2. Hydrogen and carbon monoxide are provided as synthesis gas, and at least 20% of the synthesis gas provided to the first reaction zone passes through the shear mixing device before entering the first reaction zone. 3. The method according to any one of 3. hydrogen and carbon monoxide are provided as synthesis gas, and at least a portion of the synthesis gas is introduced into the cylindrical reactor through a sparger at a height less than 50% of the reaction fluid filling height of the first reaction zone. The method according to any one of claims 1 to 4, wherein: The reactor includes a horizontally oriented ring baffle attached to an inner wall of the reactor, the ring baffle having a height that is less than 90% of the height of the liquid reaction fluid in the first reaction zone. A method according to any one of claims 1 to 5, wherein the solid part of the ring baffle extends from 5 to 30% of the diameter of the reactor. The method according to any one of claims 1 to 6, further comprising a stirrer located within the cylindrical reactor. 8. The method of claim 7, wherein the agitator and the returned reaction fluid provide a mixing energy density within the cylindrical reactor. 8. The method of claim 7, wherein the agitator is not activated. A method according to any one of claims 1 to 9, wherein the reactor is vertically oriented. A method according to any preceding claim, wherein the reactor further comprises a second reaction zone, and the reaction fluid flows from the first reaction zone to the second reaction zone without piping. 12. The method of claim 11, wherein the first reaction zone and the second reaction zone are separated by a perforated plate. 13. The method of claim 11 or 12, wherein the reactor further comprises a third reaction zone, and the reaction fluid flows from the second reaction zone to the third reaction zone without piping. 14. The method of claim 13, wherein the second reaction zone and the third reaction zone are separated by a perforated plate. 15. A method according to any preceding claim, wherein the bubbles produced by the shear mixing device have an average cell size of 10 nanometers to 3,000 microns. the reactor comprises a product outlet nozzle located in the lower part of the reactor, the reactor comprising means for preventing gas entrainment located in the bottom volume of the reactor; A method according to any one of claims 1 to 15.

Images